Reactor system

ABSTRACT

A shell and tube heat exchanger reactor with forced circulation is used to improve heat and mass transfer for exothermic liquid-liquid, gas-liquid and gas-liquid-solid reactions. Enhanced productivity and selectivity are obtained.

BACKGROUND OF THE INVENTION Description of the Prior Art

Many liquid phase oxidation and hydrogenation reactions carried out incommercial operations are highly exothermic in nature. In suchoperations, the ability to remove the heat of reaction very often limitsthe production rate obtainable for a given reactor volume. Exothermicreaction with heat removal is typically accomplished in a stirred tankreactor with a cooling jacket, a stirred tank reactor with internalcooling coils, a stirred tank reactor with an external sidestreamcooling system, or a bubble column reactor with heat transfer tubes. Inall cases, the heat of reaction is transferred from hot reaction liquidthrough a solid surface into a cooler fluid such as cooling water, arefrigerant, or to evaporate water to make steam.

Heat transfer in all of these systems is described by the followingequations

    Q=rH.sub.r V                                               (1);

    Q=UAΔT                                               (2);

and hence:

    rH.sub.r =(A/V)UΔT                                   (3);

where Q is the total heat load, r is the volumetric reaction rate, H_(r)is the heat of reaction, V is the reactor volume, U is the overall heattransfer coefficient, A is the heat transfer surface area and ΔT is thetemperature difference between the reactor liquid and the heat transferfluid. The left side of equation 3 is the volumetric heat of reaction orheat generation of the reactor, while the right side is the volumetricheat transfer capacity.

Equation 3 shows that heat generation increases with reaction rate andthat, for steady state operation, the heat transfer capacity of thesystem must be increased when the reaction rate is increased. Theequation also shows that heat transfer capacity is maximized when (1)the ratio of heat transfer area to reactor volume is maximized, (2) whenthe overall heat transfer coefficient U is maximized, and (3) when thetemperature driving force ΔT is maximized.

The area to volume ratio A/V is fixed by the geometry of the reactor andheat exchange system.

The heat transfer coefficient U is a function of fluid properties, andto a lesser extent the materials of construction of the heat exchanger.U can be increased or decreased by increasing or decreasing the flowrates of the reaction fluid and/or the heat transfer fluid. Coolingfluid flows are usually limited by pressure drop and in some cases bytemperature considerations. Depending on the heat transfer system, thereaction liquid flow rate is limited by power input to the agitator orby pressure drop considerations.

The temperature difference ΔT can be increased by increasing thereaction temperature and/or by decreasing the cooling fluid temperature.The reaction temperature is usually fixed so as to provide a givenreaction rate and/or to minimize by-product formation. Thus, it is notusually desirable to raise the reaction temperature. The temperature ofthe cooling fluid is usually limited by the temperature of availablecooling water, the cost of refrigeration or steam quality in evaporationsystems.

In conventional reactor systems for exothermic reactions, a fundamentalcharacteristic of jacketed reactor vessels is that they have a small A/Vratio. Since A increases as D, where D is the reactor diameter, and Vincreases as D², A/V decreases as the size of the reactor is increased.Thus, jacketed reactors are typically used in small volume applicationsup to 100 gallons.

Stirred tank reactors with internal cooling coils typically have ahigher ratio of A/V than jacketed vessels, particularly as the vesselsizes get larger. However, coils have several limitations. Heat transferarea is maximized by minimizing coil diameter, but pressure drop in thecoil gives a lower limit for the coil diameter. It is possible toincrease A/V by packing the reactor with coils. However, this tends tocause uneven flow distribution in the reactor which can lead to poorreactant mixing and undesirable by-product formation. It is alsomechanically difficult to support multiple coils within the reactorvessel. Reactors with internal cooling coils are thus typically used inmedium size applications of between 100 and 20,000 gallons. This reactorconfiguration is quite common in hydrogenation systems, such as inedible oil production, and in inorganic oxidations such as copperoxidation to copper sulfate.

One approach for precluding the geometric and flow constraints on theA/V ratio is to use a sidestream cooling system for external cooling.This kind of reactor configuration is often used in the oxidation ofcumene to cumene hydroperoxide for the production of phenol, and in somehydrogenation systems.

In such sidestream systems, a sidestream from the reactor is pumpedthrough a heat exchanger or other cooling system, and the cooledreaction liquor is returned to the reactor vessel. In principle, theratio of A/V is not limited by constraints associated with reactorgeometry. However, there are other potential problems associated withsuch systems. Since the cooling is accomplished outside the reactorvessel, the cooler will normally operate at a temperature that issignificantly lower than the reaction temperature. Thus, proportionallymore heat exchange area and/or coolant flow is required. Also, in gas-liquid reactor systems, such as for air oxidation, oxygen oxidation orhydrogenation reactions, the reactant gas must be prevented fromentering the external cooling system. Gas tends to disengage from theliquid and collect in pockets in high spots in the exchanger andassociated piping. This reduces the effectiveness of the heat exchanger.Gas can also collect in circulation pumps and cause cavitation or gasflooding of the pump. In oxidation systems where the gas is air oroxygen, any gas build up in the external heat exchange system couldcreate an explosion hazard. It is possible to keep the gas out of theexternal heat exchange loop. However, in many cases, this causesincreased by-product formation due to reactant starvation in reactionsystems where good mixing of reactants is important.

The A/V of bubble column reactors is typically quite high. In oneconfiguration, which is used in the production of organic acids, thebubble column is configured like a vertical shell and tube heatexchanger. The reaction takes place on the tube side while cooling fluidis circulated on the shell side. The gas is sparged into some of thetubes. Liquid circulation is caused by the gas lift effect of the gasbubbles in the sparged tubes. Thus, there is upflow and gas liquidcontacting in those tubes that are sparged. There is downflow withoutgas liquid contacting in the remainder of the tubes. This configurationhas two disadvantages. The liquid velocity on the reactant side of thetubes is limited to the bubble rise velocity which is normally between 1and 5 ft/sec. This velocity limitation limits the heat transfercoefficient U. Also, the downflow tubes are not exposed to reactant gas.This may cause a lower volumetric reaction rate and/or by-productformation as a result of gas starved conditions in the down flow tubes.

Another common bubble column configuration, which is common in theso-called Witten process for producing dimethyl terephthalate bysuccessive oxidation and esterification of p-xylene, is to use verticaltubes with reactant fluid on the outside and cooling fluid on theinside. This is mechanically very difficult to implement, but it isadvantageous if the required reaction volume is large. In such systems,feed gas is sparged into the bottom of the reactor. The gas tends tocollect into a plume such that there is a column of rising gas in onesection of the reactor and ungased downflow in the remainder. This leadsto the same conditions described above, namely limited heat transfercoefficient on the reactant side, and gas starved conditions in thedownflow regions. In addition, since the reactant side flow is notuniform, this configuration may give rise to hot spots in the vicinityof the gas plume. These hot spots can cause undesired by-productformation due to over oxidation.

Mass transfer considerations are also very important, particularly ingas liquid reaction systems. If mass transfer is reaction rate limiting,reactor productivity is determined by it. Also, reactant starvationcaused by mass transfer limitations can cause by-product formation whichlowers chemical selectivity. It is well known that these problems occurin air based chemical oxidation systems.

Air based bubble columns and stirred tank reactor systems have inherentmass transfer limitations. Oxygen mass transfer is proportional to theoxygen concentration or oxygen partial pressure in the oxygen containinggas bubble. The concentration of oxygen in an air bubble in a bubblecolumn or a stirred tank reactor is only 21% at the sparger. As oxygendissolves into the reaction liquid, where it is consumed by thereaction, and as liquid evaporates into the air bubble, the oxygenpartial pressure in the air bubbles decreases, while the partialpressure of nitrogen, which is a component of air, and the partialpressure of evaporated organic material increases. Thus, the masstransfer driving force associated with air is inherently lower than ifpure oxygen is used as the reactant gas.

In conventional stirred tank and bubble column reactor designs, theoxygen partial pressure of the exiting waste air stream must bemaintained below safety limits of 5% on an organic free basis in orderto prevent formation of flammable gas mixtures in the reactor vaporspace. Thus gas phase oxygen concentration in conventional reactordesigns is constrained between 21% at the air feed point and 5% at thewaste gas exit. In bubble columns, the air is injected at the bottom ofthe reactor. The gas bubbles rise through the liquid, and the gas phaseoxygen concentration varies from 21% at the bottom to 5% at the top ofthe reactor. In a well mixed stirred tank reactor, the average oxygenconcentration in the system is 5% throughout. Thus, for a givenoperating pressure, safety concerns in conventional reactor systemsseverely constrain the available mass transfer driving force. Thesituation can be improved somewhat by raising the overall systempressure which raises the oxygen partial pressure, or by purging theheadspace with relatively large amounts of an inert gas such asnitrogen, but these alternatives are generally very expensive.

The net result of the limited mass transfer driving force inherent inconventional air based reactor systems is that oxygen starvedconditions, and the accompanying product selectivity penalty, are morelikely to occur as reaction temperature and reactor productivity areincreased.

Another factor which limits oxygen mass transfer capacity is the degreeto which the oxygen containing gas bubbles are uniformly distributedwithin the liquid phase. If some regions of the liquid phase are notexposed to oxygen containing gas bubbles, those regions will be oxygenstarved and by-product formation will occur. Hence it is crucial to havegood gas bubble distribution throughout the reactor.

In conventional bubble column reactors or gas lift bubble columnreactors, the gas bubbles are introduced at the bottom of the reactor.They rise through the reaction liquid due to their buoyancy. The bubblescause a recirculating liquid flow pattern. In bubble column reactors,the flow tends to be up through the center of the reactor and down nearthe walls of the reactor. The oxygen containing gas bubbles tend toconcentrate in the center upflow region, which leaves the outer downflowregion gas starved and subject to by-product formation reactions. In gaslift bubble columns the oxygen is typically sparged into gas lifted heattransfer tubes such that there is liquid upflow in the tubes. Additionaltubes without spargers are provided for recirculating flow. Oxygenstarved conditions and hence by-product formation reactions prevail inthe downflow tubes.

One specific example of an oxidation system where heat transfer and masstransfer are critical is the production of aliphatic acids. Aliphaticacids are produced by liquid phase reaction of an aldehyde with oxygenaccording to the reaction:

    R--HO+1/20.sub.2 =R--OOH

The aldehydes and corresponding acids, may be linear or branched, andthe number of carbon atoms may vary from 3 to 12. The precursoraldehydes are often made using the Low Pressure Oxo (LPO) process.Hence, the derivative acids are often referred to as Oxo acids. Thealdehydes may also be obtained or produced by means other than the LPOprocess, but this class of compounds is referred to as Oxo acids, nonethe less. The source of the aldehydes is not critical to this process.

In commercial production of such acids, selectivity to acid is typicallybetween 80% and 99%. Selectivity decreases with chain length and thenumber of side chains or branches. For example, the selectivity ofpropionaldehyde which has three carbon atoms (C₃), to propionic acid, isbetter than the selectivity of valeraldehyde to valeric acid, which has5 carbon atoms (C₅); and the selectivity of valeraldehyde, which is alinear five carbon molecule, to valeric acid is higher than theselectivity of 2-methyl butyraldehyde, which is a branched C₅, to2-methyl butyric acid. In commercial practice, by-product inhibitoradditives may be added to some of these systems to improve selectivity.

In liquid phase aldehyde oxidation, the oxygen is typically introducedinto the liquid by mass transfer from gaseous air bubbles. The oxidationreactions occur in the liquid phase; either in the bulk liquid phase orin the liquid film which surrounds the air bubbles. Oxygen starvation,that is lack of dissolved oxygen in the reaction liquid, promotesby-product formation reactions and hence reduces the selectivity ofaldehyde to acid. Thus, adequate mass transfer of oxygen from the gasphase to the liquid phase is critical to maintain adequate dissolvedoxygen concentration in the liquid phase in order to suppress by-productformation reactions.

It has been found that by-product formation increases with reactiontemperature. Since reaction rate typically increases with temperature,the reaction consumes oxygen faster at higher temperature, and moreoxygen is required to prevent the onset of oxygen starved conditions.Thus, gas-liquid mass transfer limitations become worse as temperatureis increased, and. therefore, it is more difficult to prevent oxygenstarved conditions which cause by-product formation. The by-productsformed under oxygen starved conditions are formate esters, ketones andalcohols.

Since the conversion of aldehyde to acid increases with temperature, itis possible to increase reactor productivity by increasing temperature.However, if the increase in temperature moves the reaction system intothe oxygen starved regime, or makes an already oxygen starved conditionworse, by-product formation reactions increase and selectivity to aciddecreases.

Oxo acids are typically produced in air sparged stirred tank or bubblecolumn gas lift reactors. At commercial reaction conditions, theexothermic heat of reaction produced by the oxidation reactions issignificant. Although stirred tank reactors have been used for Oxo acidproduction, bubble column reactors configured as vertical shell and tubeheat exchangers are preferred because of the higher A/V ratio.

In the bubble column reactors, air is sparged into the bottom of some ofthe heat transfer tubes, while the remainder of the tubes are notsparged. This combination of sparged and unsparged tubes causes arecirculating liquid flow within the reactor. The gas causes upflow ofliquid in the sparged tubes, while downflow occurs in the remainder ofthe tubes that are not sparged. As air rises through the sparged tubes,oxygen transfers from the air into the liquid phase where it reacts withthe aldehyde to form the acid. There is no mass transfer of oxygen intothe liquid in the tubes which are not sparged.

In this reactor configuration, heat transfer occurs in all of the tubesand the ratio of A/V is high. However, the heat transfer coefficient Uis limited somewhat because the tube side flow velocity is limited tothe rise velocity of the gas bubbles which is typically between 1 and 5ft/sec. Furthermore, since a fraction of the tubes are not sparged withgas these tubes operate in the mass transfer limited or oxygen starvedmode. Thus, by-product formation is higher in the tubes which are notsparged compared to the tubes that are sparged. By-product formation isalso favored by inherent mass transfer limitations associated with usingair for the oxidant in the sparged tubes.

It will be appreciated from the above that improvements in the reactorsystem for oxidation, hydrogenation and other exothermic gas-liquidoperations are highly desired in the art. Such improvements desirablywould mitigate heat transfer limitations and improve mass transferperformance as compared to the conventional systems described above.

It is an object of the invention to provide an improved reaction systemfor oxidation, hydrogenation and other exothermic gas-liquid operations.

It is another object of the invention to provide a reactor systemcapable of mitigating heat transfer limitations and improving the masstransfer performance of exothermic gas-liquid operations.

With these and other objects in mind, the invention is hereinafterdescribed in detail, the novel features thereof being particularlypointed out in the appended claims.

SUMMARY OF THE INVENTION

Forced circulation is utilized in conjunction with a shell and tube heatexchanger reactor to improve heat and mass transfer in exothermicreactor systems. Volumetric reactor productivity and improvedselectivity are obtained thereby.

BRIEF DESCRIPTION OF THE DRAWINGS

The invention is further described in detail with respect to theaccompanying drawings in which:

FIG. 1 is a schematic side elevational view of an embodiment of thereactor system of the invention adapted for liquid-liquid reaction;

FIG. 2 is a schematic side elevational view of an embodiment of thereactor system of the invention adapted for gas-liquid reaction;

FIG. 3 is a schematic side elevational view of an embodiment of thereactor system of the invention adapted for once-through gas-liquidreaction with purge; and

FIG. 4 is a schematic side elevational view of an embodiment of thereactor system of the invention adapted for use of impeller means forenhanced gas-liquid reaction.

DETAILED DESCRIPTION OF THE INVENTION

The objects of the invention are achieved by employing a shell and tubereactor configuration such as to achieve a high heat transfer surface toreactor volume ratio A/V, together with enhanced heat transfercoefficient U, due to forced circulation of the reaction liquid. Forgas-liquid reaction systems, means are provided to achieve gascirculation throughout the entire reaction volume, thereby improvingreaction productivity and reaction selectivity. In its variousembodiments, the invention utilizes an improved reactor system that isbeneficial for conducting liquid phase reactions with two or more liquidreactants, for exothermic oxidation systems where the oxidant is eitherair or oxygen, for hydrogenation reactions and for other exothermicgas-liquid reaction systems. Such systems may or may not employ a solidcatalyst phase.

In the embodiment of the invention illustrated on FIG. 1 of thedrawings, vertical shell and tube heat exchanger reactor 1 has hollowdraft tube 2 positioned in the center thereof. Impeller means 3 arepositioned within said draft tube 2, and are adapted to recirculateliquid downward through the draft tube into bottom mixing chamber 4, andup through heat exchanger tubes 5. The reactor behaves much like a wellmixed stirred tank reactor in that bottom mixing chamber 4 provides forbulk mixing of liquid. However, because of the pumping action ofimpeller means 3, the liquid is circulated through heat exchanger tubes5 at high velocity, much like an external cooling system. Since therecirculation path is well defined and constrained by heat exchangertubes 5, the system is not subject to flow distribution problems thatcan occur when a conventional stirred tank is packed with coils. Theillustrated embodiment is adapted particularly for liquid-liquidreaction, with one liquid feed being passed through feed line 6continuing that control means 7 into upper portion 8 of reactor 1, and asecond liquid feed passing through feed line 9 having flow control means9a with said upper portion 8. Cooling water is passed to reactor 1through inlet 10, and is withdrawn through outlet 11. The liquid feed iscaused to rise upward into an upper chamber 12 in fluid communicationwith reactor 1 so as to establish liquid level 13 thereon. Productliquid is discharged from bottom mixing chamber 4 through productdischarge line 14 having control means 15. Liquid level control means 16is in communication with upper chamber 12 to receive input signal 17 asto the liquid level in the reactor and to send output signal 18 to flowcontrol means 15 so as to maintain the desired liquid level 13. Drivemotor 19 is connected to drive shaft 20, adapted to drive impeller means3. As illustrated, upper baffle means 21 and lower baffle means 22 areprovided to facilitate the desired recirculation of liquid downward inhollow draft tube 2 and upward in said tubes 5.

The illustrated system is characterized by a high A/V ratio due to itsgeometric configuration, and a high heat transfer coefficient U due tothe forced circulation flow. Thus, the FIG. 1 embodiment of theinvention is particularly suitable for exothermic liquid phasereactions.

FIG. 2 illustrates an embodiment that is suitable for exothermic gasliquid reactions that are nonflammable, particularly hydrogenationreactions, and aqueous air or oxygen based oxidation reactions. In thisembodiment, vertical shell and tube reactor heat exchanger reactor 23has hollow draft tube 24 positioned in the center thereof. Impellermeans 25 are positioned within draft tube 24, and is adapted torecirculate liquid downward through the draft tube into bottom mixingchamber 26 and up through heat exchanger tubes 27. Liquid feed is passedthrough feed line 28 containing flow control means 29 into upper portion30 of reactor 23. Reactant gas feed is passed through line 31 containingflow control means 32 and into upper chamber 33. Said control means isadapted to control reactor pressure or feed gas flow in response tosignals 32a or 32b. As will be seen from the drawing, such gasintroduction is above the level of liquid, said liquid level 34 being inreactor 23 above the positioning of impeller means 25. Cooling water isintroduced to reactor 23 through inlet 35 and is discharged throughoutlet 36. Said impeller means 25 is connected to drive shaft 37, whichis driven by drive motor 38. Upper baffle means 39 and lower bafflemeans 40 are positioned so as to facilitate the flow of liquid into thetop of hollow draft tube 24 and upward from bottom mixing chamber 26.

Reaction product is withdrawn from the bottom of reactor 23 throughproduct discharge line 41 containing flow control means 42. Liquid levelcontrol means 43 is adapted to receive an input signal 44 from reactor23 and to forward output signal 45 to flow control means 42 to controlthe liquid in reactor 23 at the desired liquid level 34. Gas iswithdrawn from upper chamber 33 through line 46 containing flow controlmeans 47 adapted for back pressure control or vent flow control asindicated by input control signals 48 and 49.

In the FIG. 2 embodiment, the reactant gas is drawn into draft tube 24through vortex action at the upper draft tube entrance under the downpumping action of impeller means 25. Thus, the impeller means creates agas dispersion within the liquid phase that is recirculated downwardthrough the draft tube into bottom mixing chamber 26 and upward throughheat transfer tubes 27. Unreacted reactant gas, inert nitrogen or byproduct gases escape into gas space 50 in upper chamber 33 above liquidlevel 34 wherein they become mixed with fresh feed gas and are drawnback into the recirculating body of liquid in reactor 23.

The reactor system of the FIG. 2 embodiment has a two fold advantageover conventional reactor systems. First, it has the beneficial fluidflow and heat transfer characteristics that are described above. Sincethe reactant gas is introduced at the top of the reactor, it is alsocirculated through the entire reactor volume, including all of the heattransfer tubes. Thus, all of the reactor volume is utilized for masstransfer, the reaction rate is maximized throughout the reactor, andby-product formation due to reactant gas starvation is minimized. In thecase of air based reactions, mass transfer can be further enhanced bythe use of oxygen feed in place of feed air.

The embodiment of the invention illustrated in FIG. 3 of the drawings isparticularly beneficial in reaction systems wherein the reactant gas canform a flammable gas mixture with the vapor above the reactant liquid,as in the air or oxygen based oxidation of organic chemicals. In suchcases, the air or other reactant gas is sparged under the liquid surfacedirectly into the impeller suction. A flammable gas mixture is formed atthe point of gas injection. However, since the gas is dispersed withinthe liquid, it is not hazardous since flame cannot propagate through theliquid. The flow path is similar to the FIGS. 1 and 2 embodiments inthat the gas liquid dispersion is pumped down through the draft tubeinto the bottom mixing chamber and up through the heat exchanger tubes.The gas then disengages from the liquid phase and collects in the gasspace above the liquid. This configuration also takes advantage ofbeneficial heat transfer and fluid flow characteristics offered by thepumped shell and tube design since the reactant gas is circulatedthroughout the entire reactor volume. The productivity of the entirereactor volume is maximized, and the potential for reactant starvedconditions that can occur in ungased tubes is minimized.

In the FIG. 3 embodiment, vertical shell and tube heat exchange reactor51 has hollow draft tube 52 positioned in the center thereof. Impellermeans 53 are positioned within said draft tube 52, preferably at theupper portion as in the other illustrated embodiments of the invention,and are adapted to recirculate liquid downward through said draft tube52 into bottom mixing chamber 54, and up through heat exchanger tubes55. Liquid feed is passed through feed line 56 containing flow controlmeans 57 preferably into upper portion 58 of reactor 51. Air or anoxygen enriched feed gas is passed through feed line 59 having flowcontrol means 59a into upper portion 58 of reactor 51, so as to be drawninto the suction of impeller means 53 along with a recirculating flow ofthe liquid in reactor 51. Cooling water is passed to reactor 51 throughinlet 60 and is withdrawn through outlet 61. The liquid is caused torise to a liquid level 63 in said upper portion 58, which is in fluidcommunication with an upper chamber 62 comprising an overhead gas phasefrom which gas is vented through gas discharge line 64 containing flowcontrol means 65. Product is discharged from bottom mixing chamber 54through line 66 containing flow control means 67. Liquid level controlmeans 68 is adapted to receive input signals 69 as to liquid level 63and to send output signal 70 to flow control means 67 so as to maintainthe desired liquid level 63. Drive motor 71 is connected to drive shaft72, adapted to drive impeller means 53. Upper baffle means 73 and lowerbaffle means 74 are provided to facilitate the desired recirculation ofliquid downward in draft tube 52 and upward in said tubes 55.

In the FIG. 3 embodiment, back pressure control means 75 are provided toreceive an input signal 76 as to the pressure in upper chamber 62 and tosend an output signal 77 to flow control means 65 in gas discharge line64. In addition, inert purge line 78 containing normal flow controlmeans, e.g. valve, 79 is used to introduce purge gas to upper chamber 62or reactor 51 above liquid level 63. Oxygen analyzer 80 is adapted toreceive input signals as to the oxygen concentration in upper chamber 62and to send output signals 82 to emergency flow control means 83 toenable additional quantities of inert purge gas to flow throughemergency flow line 84 to reactor 51 or upper chamber 62 above liquidlevel 63.

In flammable systems, the potential to form flammable gas mixtures inthe waste gas stream must be eliminated. For example, in the oxidationof an organic liquid with air, the oxygen content in the waste gas mustbe reduced below the flammable oxygen concentration which is typicallybetween 8% and 12%. In practice, the oxygen concentration is reduced tobelow 5% to provide an adequate safety margin. When air is used in thisembodiment, the oxygen concentration in the gas can be reduced byreaction from 21% at the point of injection to less than 5% in the wastegas. Conventional reactors are operated in this way. Alternatively,nitrogen or other diluent gas can be added to the waste gas to reducethe oxygen concentration to less than 5%. If pure or nearly pure oxygenis used in this embodiment, the oxygen must also be reacted away or aninert diluent is added to the waste gas, as shown above, but the masstransfer performance of the system is improved due to the higher oxygenconcentration.

FIG. 4 of the drawings shows the preferred embodiment for use in systemswhere the reactant gas can form a flammable gas mixture with the vaporabove the liquid phase. This embodiment is particularly beneficial inthe oxidation of organic chemicals with pure oxygen. In this embodiment,the oxygen is injected below the liquid surface and the flow pattern isthe same as that described in the embodiment above. However, in thisembodiment, a gas containment baffle is used to direct the gas flow fromthe top of the heat exchanger tubes back into the draft tube suctionzone. Nitrogen or other inert gas is passed through the gas space abovethe reactor to insure that any oxygen that escapes from under the gascontainment baffle is diluted to a concentration below 5%. Thisarrangement is a novel and advantageous modification of a conventionalLOR reactor system having enhanced heat transfer capability to make theLOR system suitable for highly exothermic reactions.

In the FIG. 4 embodiment, vertical shell and tube heat exchanger reactor101 has hollow draft tube 102 positioned in the center thereof. Impellermeans 103 are positioned within said draft tube 102, and are adapted torecirculate liquid reactant downward through the draft tube into bottommixing chamber 104, and upward through heat exchanger tubes 105. Anorganic liquid feed is passed through feed line 106 containing flowcontrol means 107 into upper portion 108 of reactor 101. An oxygen feedline 109 having flow control means 109a is used to pass oxygen or anoxygen-containing gas into said upper portion 108. Cooling water ispassed to reactor 101 through inlet 110, and is withdrawn through outlet111. The organic liquid feed is caused to rise upward into upper chamber112 so as to establish liquid level 113 therein. Product liquid isdischarged from bottom mixing chamber 104 through product discharge line114 having control means 115 therein. Liquid level control means 116 isin communication with measuring means at upper chamber 112 to receiveinput signal 117 as to the liquid level in the reactor, and to sendoutput signal 118 to flow control means 115 so as to maintain thedesired liquid level 113. Drive motor 119 is connected to drive shaft120, adapted to drive impeller means 103. Upper baffle means 121 andlower baffle means 122 are provided to facilitate the desiredrecirculation of liquid downward in hollow draft tube 102 and upward insaid tubes 105. It should be noted that gas containment baffle 123 ispositioned in the upper portion of reactor 101 above upper baffle 121.While maintaining fluid communication between the liquid in reactor 101and the liquid in upper chamber 112, said gas containment baffle 123serves to minimize undesired flow of unreacted gas into said upperchamber 112 to liquid level 113 and the overhead gas phase.

In the FIG. 4 embodiment, back pressure control means are provided toreceive an input signal 125 as to the pressure in the gas phase 126 inupper chamber 112 and to send output signal 127 to flow control means128 in gas discharge line 129, which contains condenser 130 to whichcooling water is added through line 131 and from which cooling water iswithdrawn through line 132. In addition, nitrogen or other inert gaspurge line 133 is used to introduce purge gas to gas phase 126 in upperchamber 112. Oxygen analyzer 134 is adapted to receive input signals asto the oxygen concentration in said gas phase 126 and to send outputsignals 136 to normal purge flow control means 137 and to send outputsignals 138 to emergency flow control means 139 to enable additionalquantities of purge gas to flow to said gas phase 136 in upper chamber112, if needed.

In the production of Oxo acids, i.e. aliphatic acids produced by theoxidation of the corresponding aldehydes, much higher liquid circulationvelocities, and therefore much higher heat transfer coefficients, areachieved using the reactor system of the invention relative to gas liftshell and tube bubble column reactors.

Oxygen was used in the reactor system of the invention to improveselectivity in the oxidation of Oxo aldehydes to the correspondingacids. With oxygen, the partial pressure of oxygen in the oxygencontaining gas bubbles within the oxidation reactor is significantlygreater than the inherently limited oxygen partial pressure in air.Consequently, the driving force for mass transfer is greater, and thelikelihood of oxygen starved conditions which cause by-product formationis lower, with oxygen.

The subject reactor of the invention is a well mixed stirred tankreactor system, consequently oxygen bubbles are uniformly distributedthroughout the liquid. Thus, with said reactor there are no zones thatare oxygen starved due to poor gas liquid contacting. Also, with saidreactor, the requirement that the gas bubbles have a concentration of 5%or less does not apply. Consequently, depending on the vapor ressure ofthe liquid which acts as a diluent, the average oxygen concentration inthe gas bubbles is much higher than it is in a conventional reactor withair. In systems with a very low liquid vapor pressure, the averageoxygen concentration can approach 95% or higher. This compares favorablyto the average 5% oxygen concentration in a conventional air basedstirred tank reactor and to the average of 13% in a bubble columnreactor.

The overall higher mass transfer rate gives rise to improved oxygen masstransfer which increases the amount of oxygen available for reaction inthe liquid phase and thereby reduces selectivity losses that areassociated with oxygen starved conditions. The overall higher masstransfer rate also allows for operation at lower temperature andpressure than in conventional air based reaction systems. In particular,the operation at lower temperature further reduces by-product formationand increases selectivity.

A preferred practice of the invention is to use a shell and tube heatexchanger with a draft tube in the center as shown in FIG. 4. The heatload on the system is determined from the heat of reaction, the desiredvolumetric reaction rate and the reactor volume as shown in Equation 1.Fixing the reaction rate generally fixes the reaction temperature andpressure as well. Once the heat load, the reaction temperature andpressure are known, standard methods are used to determine U the overallheat transfer coefficient, A the required heat transfer area, and ΔT thetemperature difference. These procedures are described below.

Once the reactor volume is determined, a mixing impeller/pump is chosenbased on power input criteria which is required for adequate reactantmixing, and flow criteria for liquid and gas circulation. These criteriaare well known to those who routinely design mixing and heat exchangeequipment. For gas liquid mixing systems, the power required is usuallyabout 5 HP/1000 gallons of liquid, but this number can vary considerablydepending on the reaction system. The flow criteria are two fold. First,it is desirable to maximize flow velocity through the heat transfertubes in order to maximize the heat transfer coefficient U. However,pressure drop through the tubes increases as the velocity squared. Thus,there is an optimum velocity for a given system. The second pumpingcriteria, which is important in gas liquid systems, is that the liquidvelocity within the draft tube must be maintained above a minimum of 1ft/s, but preferably above 2.5 ft/s in order to insure that the gas isdrawn downward through the draft tube.

Since the reactor design is based upon the required mixing impeller/pumpcharacteristics, the size and speed of the impeller are chosen on thebasis of the required mixing power and reactor volume. This fixes thepumping characteristics for the impeller including the flow versus headcurve. The heat transfer area A is estimated from Equation 2 using anestimated U. Once A and the flow are estimated, the geometry of thesystem can readily be determined.

The design objective is to maximize the ratio of A/V and therebymaximize the volumetric heat transfer capacity, which is turn maximizesreactor productivity. Thus small diameter heat exchanger tubes aredesired. However, as tube size decreases, pressure drop within the tubesincreases. Thus there is an optimum for every design case. Using theestimated flow, the heat exchanger tube diameter, number of tubes andtube length are varied to give the required A and acceptable values forpressure drop. Usually 1" or 3/4" diameter heat transfer tubes areoptimum.

Once the tube diameter, length and flow are known, detailed calculationsfor U and ΔT can be completed. The size and speed of the mixing impellermay have to be increased if the calculated U is not large enough tosatisfy Equation 1. Alternatively, the heat transfer area A can beincreased. Those skilled in the art of heat transfer will know how toadjust the design parameters to satisfy the heat transfer requirements.It should be noted that in the final design both heat transfer andmixing power criteria must be satisfied.

Once the impeller size is determined, and the diameter, length andnumber of tubes is fixed, the remainder of the reactor geometry can becompleted. The draft tube diameter is fixed by the diameter of themixing impeller. The heat transfer tubes are arranged in a standardtriangular pitch around the draft tube. A triangular pitch is preferredbecause it gives a higher A/V than either a square pitch or a radiallayout. Once the tubes are arranged the overall diameter of the reactorvessel is fixed.

The upper head geometry depends on which reactor configuration is used.In the liquid reactor configuration, and the configuration for gasliquid reactions (known as the AGR configuration), where gas is drawndown from the gas space, the main design constraint is to insure thatthe flow pattern across the top of the heat exchanger tubes is symmetricso as to provide an even flow distribution in the tubes. This isaccomplished by using a conical outer shell in conjunction with a flaredconical draft tube inlet section. This symmetric conical arrangementinsures that the flow exiting the tubes accelerates gradually anduniformly such that the impeller suction draws from all of the tubes atan equal rate.

In the cases where the reactant gas is injected below the liquidsurface, passes once through the reactor and the unreacted gas is ventedas waste gas, the waste gas must be prevented from being entrained inthe draft tube suction flow. The symmetric conical arrangement describedabove is also necessary. In this case, the exit of the symmetric conearrangement must be positioned such that the flow path from the top ofthe cone to the draft tube suction is long enough for the bubbles todisengage from the liquid before the liquid enters the draft tube.

In cases where a gas containment baffle is used in the upper head, thesymmetric conical arrangement is also required. Other design constraintsrelated to the gas containment baffle are disclosed in the KingsleyPatent, U.S. Pat. No. 5,451,349.

The bottom mixing chamber is made from a conical or dished head of thesame diameter as the heat exchanger tubesheet. A conical head is shownin the drawings. The volume of this head can be adjusted so the totalreactor volume matches the desired reactor volume.

A cross baffle is used in the lower head to help insure even flowdistribution across the heat exchanger tubes. The cross baffle serves tosegment the discharge flow from the bottom of the draft tube into fourequal parts which are directed in the radial direction.

In gas liquid reactor systems, where product is withdrawn continuously,an additional baffle is used to separate gas bubbles from the liquidproduct. This is not a critical part of this invention since there aremany ways to accomplish this.

If the reactor system is used in a gas liquid reaction where theunreacted gas is directed into a gas space above the liquid as shown inFIGS. 2-4, the mixing impeller is positioned in the top of the drafttube so as to (1) induce the vortex action to draw gas down into theliquid phase, or (2) disperse gas which is fed under the liquid surfacesuch as when the reactant gas forms a flammable vapor mixture with theliquid. In both cases putting the pump near the top of the draft tubeand introducing the gas near the top of the draft tube insure thatreactant gas is circulated throughout the reactor volume. Gas could beintroduced into the bottom of the reactor in these cases but the gaswould follow the upflow through the heat exchanger tubes and exit intothe gas space above. The draft tube would remain ungased which woulddefeat one of the advantages of this system, namely the uniformdistribution of gas throughout the reaction volume.

In the case of liquid reactions as shown in FIG. 1, and in the casewhere a gas containment baffle is used as shown in FIG. 4, the mixingimpeller/pump can be placed in either the top of the draft tube or thebottom. In the liquid reactor, the primary function of the impeller isto pump liquid so its position is not critical. In the case of the gasliquid reactor system with a gas containment baffle, the gas containmentbaffle serves to direct unreacted gas into the draft tube where the gasis drawn downward by the liquid flow. Hence in this case, putting theimpeller/pump in the bottom of the reactor does not prevent gas frombeing distributed uniformly throughout the reactor.

The mixing impeller/pump can be any axial flow device such as a marinepropeller or a fluid foil impeller such as the Lightnin A-315. Thepreferred embodiment for gas liquid reactor systems is the doublehelical impeller described by Litz et al. in U.S. Pat. No. 4,900,480.

The advantages of the system described in FIG. 4 were demonstrated in a280 gallon heat exchanger reactor system used in the oxidation of2-ethylhexaldehyde to 2-ethyhexanoic acid. oxygen was sparged into thetop of the draft tube where it was dispersed by the helical mixingimpeller and pumped downward through the draft tube to the bottom of thereactor. The gas liquid mixture then passed upward through the heatexchanger tubes into the upper head. The conical upper head and gascontainment baffle served to direct the unreacted oxygen bubbles backinto the draft tube where they were redispersed and recirculated by themixing impeller. The aldehyde was fed continuously into the upper headof the reactor. The product acid was withdrawn continuously from thelower head.

The reactor was operated in the manner of an LOR (Liquid OrganicReactor) system as in said Litz et al, U.S. Pat. No. 4,900,480disclosure. The gas space of the reactor above the gas containmentbaffle was continuously inerted with nitrogen to maintain the oxygenconcentration in the gas space in a safe range.

In the test, the 280 gallon demonstration reactor was run in parallelwith a train of air sparged gas lift shell and tube bubble columnreactors operated in series. The advantages of the new reactor systemover the existing technology are shown in the Table below. Thedemonstration reactor was run at the same volumetric reaction rate asthe average of the air based reactors. The reaction efficiency isdefined as the product of the fractional conversion of aldehyde and theselectivity of converted aldehyde to acid. The temperature and pressurevariations are given as a range because the conventional reactorsoperate at different temperatures and pressures. The highest temperatureis maintained in the reactor with the highest reaction rate in order tomaintain the reaction rate and maintain a high ΔT driving force for heattransfer.

                  TABLE                                                           ______________________________________                                        Reaction Rate       Equivalent                                                Reaction Efficiency 2.5% Higher than Air                                                          Reactors                                                  Temperature         8° C. to 43° C. Lower                                           than Air Reactors                                         Pressure            30 psig to 50 psig                                                            Lower than Air                                                                Reactors                                                  Volume Adjusted Waste                                                                             36% of Air Reactor                                        Gas Flowrate        Total                                                     ______________________________________                                    

At equivalent reaction rate, the reactor system of the invention gave anoverall reaction efficiency which was 2.5% higher in a single reactorthan was achieved in a train of three conventional reactors operated inseries. Furthermore, the operating conditions were much less severe. Thesubject reactor operated at from 8° C. to as much as 43° C. lower intemperature, and from 30 psig to 50 psig lower in pressure than theconventional reactors. The lower temperature operation was madepractical by the improved heat transfer performance of the system. Thetemperature driving force ΔT can be reduced as U is increased due tohigher tube side flow velocities resulting from forced circulation.Also, since the reactor was configured as an oxygen based LOR the flowrate of waste gas was reduced to only 36% of the air reactors on areactor volume equivalent basis.

From the above, it will be appreciated that the reactor system of theinvention is a highly desirable advance in the art, enabling enhancedheat and mass transfer to be achieved in exothermic reactor systems. Thevolumetric reactor productivity can thus be maximized. Selectivitylosses associated with oxygen starved conditions are also desirablyreduced. The overall high mass transfer rate, which enables operation atlower temperature and pressure, advantageously reduces by-productformation and thus increases selectivity. The reactor system of theinvention provides an improvement desired in the art for liquid- liquid,gas-liquid and gas-liquid-solid reaction systems that are heat and masstransfer limited, enhancing productivity and selectivity performance ofoperations carried out therein.

I claim:
 1. A reactor system for the carrying out of exothermicreactions of a liquid first reactant and one of a gas and liquid secondreactant comprising:(a) a vertically positioned tube and shell reactorvessel having a hollow draft tube in the center thereof and heatexchanger tubes in the annular space between the hollow draft tube andthe outer wall of the reactor vessel, said reactor vessel having anupper space above and a hollow mixing chamber below said hollow drafttube and said heat exchanger tube; (b) impeller means positioned withinsaid hollow draft tube to cause the flow of the liquid first reactantdownward through the hollow draft tube into the bottom mixing chamberand upward through said heat exchanger tubes as a substantially uniformdispersion of reactants and into said upper space in the reactor vessel;(c) an upper chamber positioned above and in fluid communication withsaid reactor vessel; (d) conduit means for introducing said liquid firstreactant into the reactor vessel and for introducing said one of saidgas and liquid second reactant into one of said reactor vessel and saidupper chamber for recirculation with the liquid first reactant downwardthrough the hollow draft tubes into the bottom mixing chamber and upwardthrough said heat exchanger tubes into said upper space; (e) conduitmeans for withdrawing product liquid from the reactor vessel; (f)conduit means for flowing cooling fluid to the reactor vessel for theremoval of exothermic heat of reaction generated within said reactorvessel; (g) control means for maintaining a desired liquid level withinone of said reactor vessel and said upper chamber, and (h) cross bafflemeans located at the lower end of said hollow draft tube, and having adesign such that there is even flow distribution of recirculatingreactants across said heat exchanger tubes.
 2. The reactor system ofclaim 1 and including baffle means at the upper end of said hollow drafttube.
 3. The reactor system of claim 1 in which said control meansmaintain a desired liquid level in said upper chamber, said conduitmeans for introducing one of said gas and liquid second reactantintroduce a liquid second reactant into the reactor vessel.
 4. Thereactor system of claim 3 and including baffle means at the upper ofsaid hollow draft tube.
 5. The reactor system of claim 1 in which saidconduit means for introducing said one of said gas and liquid saidsecond reactant are gas conduit means for introducing a gas secondreactant into the upper chamber above said liquid level in said one ofsaid upper chamber and upper space in the reactor vessel.
 6. The reactorsystem of claim 5 and including baffle means at the upper end of saidhollow draft tube.
 7. The reactor system of claim 1 in which saidconduit means for introducing said one of said gas and liquid secondreactant introduce one of air, an oxygen containing feed gas andhydrogen into the reactor vessel below said liquid level in one of saidupper chamber and upper space in the reactor vessel.
 8. The reactorsystem of claim 7 and including conduit means for passing inert purgegas through said one of said upper chamber and upper space in thereactor vessel above said liquid level.
 9. The reactor system of claim 8and including baffle means at the upper end of said hollow draft tube.10. The reactor system of claim 9 and including control means to passadditional quantities of inert purge gas to said one of said upperchamber and upper space in the reactor vessel above said liquid levelfor emergency purposes.
 11. The reactor system of claim 1 in which saidcontrol means maintain a liquid level in said upper chamber, saidconduit means for introducing said one of said gas and liquid secondreactant introduce an oxygen-containing gas second reactant into saidone of said upper chamber and upper space of the reactor vessel belowsaid liquid level, the conduit means for introducing said liquid firstreactant in the reactor vessel introduce an organic liquid as saidliquid first reactant.
 12. The reactor system of claim 11 and includinga gas containment baffle positioned in said one of said upper chamberand upper space of the reactor vessel, above said impeller means, saidgas containment baffle minimizes the passage of undissolved gas bubblesupward to the liquid level in the upper chamber, and to the overhead gasin said upper chamber above the liquid level.
 13. The reactor system ofclaim 12 and including baffle means at the upper end of said hollowdraft tube.
 14. The reactor system of claim 11 and including conduitmeans for passing inert purge gas through said one of said upper chamberand upper space in the reactor vessel above said liquid level.
 15. Thereactor system of claim 14 and including control means to passadditional quantities of inert purge gas to said one of said upperchamber and upper space in the reactor vessel above said liquid levelfor emergency purposes.
 16. The reactor system of claim 1 in which saidhollow draft tube has a conically flared portion at the upper endthereof.